Asphaltene hydrodesulfurization with small catalyst particles disposed in a guard chamber-main reactor system

ABSTRACT

A HYDRODESULFURIZATION OF A CRUDE OIL OR A REDUCED CRUDE CONTAINING THE ASPHALTENE FRACTION PROCEEDS AT UNEXPECTEDLY LOW TEMPERATURES BY UTILIZING A SUPPORTED GROUP VI AND GROUP VIII METAL-CONTAINING HYDRODESULFURIZATION CATALYST WHEN THE CATALYST PARTICLES ARE VERY SMALL AND HAVE A DIAMETER BETWEEN ABOUT 1/20 AND 1/40 INCH AND WHEN THE CATALYST IS DISPOSED IN A GUARD CHAMBER-MAIN REACTOR SYSTEM. THE APPORTIONING OF THE TOTAL CATALYST BETWEEN THE GUARD CHAMBER AND THE MAIN REACTOR AND THE FREQUENCY OF RENEWEL OF THE GUARD CHAMBER ARE ESTABLISHED ACCORDING TO THE DISCLOSED EQUATIONS TO MINIMIZE THE PRESSURE DROP ACROSS THE CATALYST BY MAINTAINING ON STREAM AT A GIVEN TIME A MINIMUM PROPORTION OF THE TOTAL CATALYST. THE EQUATIONS HAVE PARTICULAR APPLICABILITY TO THE HYDRODESULFURIZATION PROCESS BUT THEY CAN ALSO BE APPLIED TO OTHER PROCESSES.

1971 M. D. FRASER ET AL 3,563,887

ASPHALTENE HYDRODESULFURIZATION WITH SMALL CATALYST PARTICLES DISPOSED IN A GUARD CHAMBER-MAIN REACTOR SYSTEM Filed Oct. 25, 1968 7 Sheets-Sheet 1 rfupzmruei r PRODUCE 00 5 a-wr I L FUR //v PROOUCT 355/0041. R

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ASPHALTENE HYDRODESULFURIZATION WITH SMALL CATALYST PARTICLES DISPOSED IN A GUARD CHAMBER-MAIN REACTOR SYSTEM Filed Oct. 25, 1968 7 Sheets-Sheet 2 g 750 K m a EAS/NG 9 PRESSURE M x Q Q 700 i3 a/wsrm r {a E HYDROGEN pesssues 1 w an? A Q Q q 0102030405060 70 90/00 0:47:41. v.57 AGE 0/045 7 Pi 6 Q 6 S 95 I} g 90 fifi a? Q g 750 770 a 702 025 x AVERAGE 051cm TEMPERATURE: F.

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ALLEN E. SOMERS 1971 FRASER ET AL A'SPHALTENE HYDRODESULFURIZAI'ION WITH SMALL CATALYST PARTICLES A I DISPOSED IN A GUARD CHAMBER-MAIN REACTOR SYSTEM Filed 001.. 25, 1968 r 7 Sheets-Sheet A Feb. 16, M D FRASER ET AL ASPHALTENE HYDRODESULFURIZATION WITH SMALL CATALYST PARTICLES DISPOSED IN A GUARD CHAMBER-MAIN REACTOR SYSTEM Filed Oct. 25, 1968 7 Sheets-Sheet 5 M sw/Na M #54001 I W FED I v p ucr A MAI/V REACTOR Rap M SWING A, i 254cm 2 TIME //v .SWl/VG eucroe X) AMOUNT or emu Ysr /A/ sw/A/a 254cm 1 N VENTORS MALCOLM D. FRASER ALLEN E. SOMERS Feb. 16, 1.971

, M. D. FRASER ET AL 3,563,887 ASPHALTENE HYDRODESULFURIZA'TION WITH SMALL CATALYST PARTICLES UARD OR SYSTEM DISPOSED IN A G CHAMBER-MAIN REACT Flled Oct 25, 1968 I 7 Sheets-Sheet 6 5470 r/a/v o E P E m m) m www M FR w Ra & WE F a .5 7% m T ML MC 5 my 2% w 7 W0 {T E. 0 0 0 w. m \Qmvc q T MWWWRMXQRNYU K. 33. vQwn wk IN VENTORS C4 724 YST 465- DA vs MALCOLM D. FRASER ALLEN E. SOMERS United States Patent Other:

U.S. Cl. 208-216 13 Claims ABSTRACT OF THE DISCLOSURE The hydrodesulfurization of a crude oil or a reduced crude containing the asphaltene fraction proceeds at unexpectedly low temperatures by utilizing a supported Group VI and Group VIII metal-containing hydrodesulfurization catalyst when the catalyst particles are very small and have a diameter between about and 4 inch and when the catalyst is disposed in a guard chamber-main reactor system. The apportioning of the total catalyst between the guard chamber and the main reactor and the frequency of renewal of the guard chamber are established according to the disclosed equations to minimize the pressure drop across the catalyst by maintaining on stream at a given time a minimum proportion of the total catalyst. The equations have particular applicability to the hydrodesulfurization process but they can also be applied to other processes.

The present invention relates to a process for the hydrodesulfurization of a crude oil or a reduced crude oil in the presence of a supported Group VI and Group VIII metal hydrodesulfurization catalyst having an exceptionally small particle size. Substantially all or a large proportion of the catalyst particles of the present invention have a diameter of between about & and inch.

Although nickel-cobalt-molybdenum is the preferred active metals combination for the catalyst of this invention, other combinations can be utilized such as cobalt-molybdenum, nickel-tungsten and nickel-molybdenum. Alumina is the preferred supporting material but other non-cracking supports can also be used such as silica alumina and silica magnesia.

Hydrodesulfurization catalysts comprising supported Group VI and Group VIII metals, such as nickel-cobaltmolybdenum on alumina, having a particle size as small as the catalyst particles of the present invention were not heretofore considered advantageous for use in a large or commercial scale because a bed comprising particles of the small size of the present invention induces an extremely high pressure drop, which is highly deleterious to a hydrodesulfurization process which has a limited inlet pressure because the temperature required by a catalyst to accomplish a given degree of desulfurization increases with loss of hydrogen pressure.

The present invention relates to a hydrodesulfurization process in which the small particle size catalyst is utilized in a manner which manifests an unexpectedly high activity so that hydrodesulfurization of crude oil charge to any desired sulfur level, such as a 1 percent sulfur level, proceeds at an unexpectedly low temperature. Although extrapolation of the initial temperature required to produce a liquid product having a 1 percent sulfur content with 5; inch diameter and inch diameter NiCoMo catalyst particles, which are abo e the size of the present invention, indicates that the temperature requirement would be lower with the small catalyst particles of this invention we have found that the small size 3,563,887 Patented Feb. 16, 1971 NiCoMo catalyst particles of the present invention permit the use of a hydrodesulfurization temperature which is considerably lower than the temperature which would be expected by extrapolation of the temperature data obtained with larger size catalyst particles. Moreover, the very discovery that hydrodesulfurization with the catalyst of the present invention could be carried out at an unexpectedly low temperature has heretobefore been obscured by the extremely high pressure drop through a bed of the small size catalyst particles of the present invention. The reason is that in a hydrodesulfurization process pressure drop itself increases the temperature requirement to achieve a given degree of desulfurization usually by an extent which equals or exceeds the temperature advantage due to the small particle size of this invention.

It is seen that there are two unexpected features surrounding the present invention. The first is that there is an unexpectedly great temperature advantage achievable in a hydrodesulfurization process by employing a bed of catalyst having a particle size of the present invention. FIG. 1 (all figures are discussed in detail herein below) shows that the hydrodesulfurization temperature H required to produce a residual product having 1 percent sulfur with a inch catalyst of the present invention is much lower than what would be expected by extrapolation of the line connecting the data obtained with and inch catalyst particles even though the surface area defined by the pores of all three catalysts is about the same. The second feature is that the unexpected temperature advantage is completely disguised by the ordinary approach to its determination, i.e. by making a test in a reactor with relatively large size catalyst particles as a blank and then making a test in the same reactor under the same conditions except that the catalyst particle size is within the range of the present invention (particle size being the only variable changed in the two tests). In this regard, the vertical dashed line in FIG. 2 shows that if a inch catalyst, which is larger than a catalyst of this invention, is tested in a foot diameter reactor and then a inch catalyst of this invention is tested in the same reactor under unchanged conditions, including an unchanged space velocity, the pressure drop in the inch catalyst bed in the same reactor is so much greater than that for the inch catalyst that the pressure drop increase iself would easily nullify the temperature advantage achievable because of particle size and therefore the advantage of the present invention would be completely masked. The horizontal dashed line of FIG. 2 shows that if the inch catalyst is tested in a 9.5 diameter reactor a comparable pressure drop can only be achieved if the inch catalyst is tested in an 11 foot diameter reactor, when both tests are performed at a liquid hourly space velocity of one. Therefore, it is only by making the tests in two different reactors, to equalize pressure drop, that the temperature advantage of the inch catalyst becomes apparent. It is clear that not merely one but rather two variables must be changed to reveal the advantage of the present invention.

The great effect of pressure drop upon temperature requirements to produce a hydrocarbon product having a 1 percent sulfur content is demonstrated by reference to FIG. 3. In FIG. 3 the solid line represents a hydrodesulfurization process having a constant hydrogen partial pressure of 18304850 p.s.i.a. The dashed line represents a decreasing hydrogen partial pressure starting with the 1830-1850 p.s.i.a. range to a range of 1720-1740 p.s.i.a. which reduction is caused by the recycle hydrogen stream becoming progressively diluted with other gases. FIG. 3 shows that as the hydrogen partial pressure progressively decreases the temperature required to produce a 1 3 percent sulfur product progressively increases, so that it is considerably above the temperature required with con stant hydrogen partial pressure. Since pressure drop due to flow across a catalyst bed similarly reduces partial hydrogen pressure, FIG. 3 illustrates the detrimental effect upon reaction temperature of pressure drop through a catalyst bed of this invention.

The charge to the process of this invention can be a full crude or a reduced crude containing substantially all of the residual asphaltenes of the full crude. The residual asphaltenes are deficient in hydrogen and comprise only about percent of the charge oil but contain substantially all of the metallic components present in the crude, such as nickel and vanadium. Since the desulfurization catalyst has a greater activity for demetalization than for desulfurization, it removes nickel and vanadium from a charge stock more rapidly than it removes sulfur. These metals deposit most heavily on the outermost regions of the catalyst cross section and tend to reduce the desulfurization activity of the catalyst. Nickel and vanadium removal constitutes substantially the entire deactivation of the catalyst while sulfur and nitrogen removal contributes very little to catalyst deactivation. Furthermore, the asphaltenes comprise the highest boiling fraction of the full crude and contain the largest molecules in the crude. These large molecules are the ones least able to penetrate catalyst pores and most likely to plug these pores. The present invention is directed towards the hydrodesulfurization of a full crude or a residual oil containing substantially the entire asphaltene fraction of the crude from which it is derived and which therefore contains 95 to 99 weight percent or more of the nickel and vanadium content of the full crude. The nickel, vanadium and sulfur content of the liquid charge can vary over a wide range. For example, nickel and vanadium can comprise 0.002 to 0.03 weight percent (20 to 300 parts per million) or more of the charge oil while sulfur can comprise about 2 to 6 weight percent or more of the charge oil. If an oil containing smaller quantities of nickel, vanadium and sulfur is processed, such as a furnace oil, considerably lower temperature conditions, pressures as low as 1000 pounds per square inch, lower gas circulation rates and hydrogen of lower purity than required by this invention, will suffice to produce a liquid product containing 1 percent sulfur, and therefore the process of the present invention will not be essential.

As the hydrodesulfurization reaction proceeds, nickel and vanadium removal from the charge oil tends to occur preferentially over sulfur removal. However, deposition of nickel and vanadium upon the catalyst results in a greater degree of catalyst deactivation than does sulfur removal because the removed metals deposit upon the catalyst whereas sulfur removed from the charge escapes as hydrogen sulfide gas. Low hydrodesulfurization temperatures tend to inhibit metal removal from the charge and thereby reduce catalyst deactivation. Since the hydrodesulfurization reaction is exothermic, it is important to quench the reactor to maintain a reaction temperature as low as the small catalyst size of this invention permits to obtain the desired degree of desulfurization in order to inhibit catalyst deactivation. Unnecessarily high temperatures by encouraging catalyst deactivation will result in loss of the initial temperature advantage of the catalyst of this invention. Quenching is advantageously accomplished by dividing the total catalyst bed into a plurality of relatively small beds in series and injecting relatively cool hydrogen between the beds, as demonstrated below. It is seen that there is a high degree of interdependence between the use of a high metals content asphaltene charge, the small size catalyst particles of this invention, and the use of a quench to insure that the reactor remains at a temperature as low as the catalyst size permits.

The hydrodesulfurization process of this invention employs conventional reaction conditions such as, for eX- ample, a hydrogen partial pressure of 1,000 to 5,000

pounds per square inch, generally, 1,000 to 3,000 pounds per square inch, preferably, and 1,500 to 2,500 pounds per square inch most preferably. Reactor design limitations usually restrict inlet pressures under the conditions of the present invention to not more than 2,000, 2,500, or 3,000 p.s.i.g. It is the partial pressure of hydrogen rather than total reactor which determines hydrodesulfurization activity. Therefore, the hydrogen stream should be as free of other gases as possible. Furthermore, since reactor design limitations restrict hydrogen inlet pressures, hydrogen pressure drop in the reactor should be held as low as possible.

The gas circulation rate can be between about 2,000 and 20,000 standard cubic feet per barrel, generally, or preferably about 3,000 to 10,000 standard cubic feet per barrel of gas preferably containing 85 percent or more of hydrogen. The mol ratio of hydrogen to oil can be between about 8:1 and 80:1. Reactor temperatures can range between about 650 and 900 F., generally, and between about 680 and 800 F., preferably. The temperature should be low enough so that not more than about 10, or percent of the charge will be cracked to furnace oil or lighter. At temperatures approaching 800 F. the steel of the reactor walls rapidly loses strength and unless reactor wall thicknesses of 7 to 10 inches or more are utilized a temperature of about 800 F. constitutes a metallurgical limitation. The liquid hourly space velocity in each reactor of this invention can be between about 0.2 and 10, generally, between about 0.3 and 1 or 1.25, preferably, or between about 0.5 and 0.6 most preferably.

The catalyst employed in the process is conventional and comprises sulfided Group VI and Group VIII metals on a support such as nickel-cobalt-molybdenum or cobaltmolybdenum on alumina. Hydrodesulfurization catalyst compositions suitable for use in the present invention are described in U.S. 2,880,171 and also in U.S. 3,383,301. However, an essential feature of the catalyst particles of the present invention is that the smallest diameter of these particles is considerably smaller than the diameter of hydrodesulfurization catalyst particles of the prior art. The smallest diameter of the catalyst particles of the present invention is broadly between about V and 4 inch, preferentially between and 1 inch, and most preferably between about and inch. Particle sizes below the range of this invention would induce a pressure drop which is too great to make them practical. The catalyst can be prepared so that nearly all or at least about 92 or 96 percent of the particles are within the range of this invention. The catalyst can be in any suitable configuration in which the smallest particle diameter is within the range of the present invention such as roughly cubical, needle-shaped or round granules, spheres, cylindricallyshaped extrudates, etc. By smallest particle diameter we mean the smallest surface to surface dimension through the center or axis of the catalyst particle, regardless of the shape of the particle. The cylindrical extrudate form having a length between about V and A inch is highly suitable.

Since the asphaltene molecules which are hydrodesulfurized in accordance with the present invention are large molecules and must enter and leave the pores of the catalyst without plugging the pores, in order to obtain good aging properties most of the pore volume of the catalyst of this invention should be in pores above A. in size. Advantageously to percent of more of the pore volume should be in pores of 50 A. or more. Most preferably, to percent or more of the pore volume should be in pores above 50 A. in size. Catalysts having smaller size pores have good initial activity but poor aging characteristics due to gradual plugging of the pores by the alphaltene molecules. For example, catalyst A below exhibited good activity in the process of this invention for about one month while catalyst B below exhibited good activity for about three months.

As indicated above, as the diameter of conventional hydrodesulfurization catalyst particles progressively decreases within a range which is above the range of the present invention, hydrodesulfurization of a crude oil to a one percent sulfur level proceeds at progressively lower temperatures. However, the following tests show that the diminishing of catalyst diameter size to a level within the range of the present invention results in an unexpectedly great reduction in hydrodesulfurization temperature which is much greater than indicated by the particle diameterternperature relationship exhibited by larger size particles. However, counteracting this temperature advantage is the fact that the small catalyst particle diameters of the present invention result in a large pressure drop through a catalyst bed comprising them, and this pressure drop tends to nullify the temperature advantage achievable with the catalyst of the present invention because hydrodesulfurization temperature requirements increase as hydrogen partial pressures decrease.

In accordance with the present invention the pressure drop in a system comprising fixed beds of the small size catalyst particles is reduced or minimized by an arrangement wherein only a portion of the total catalyst employed is on-stream at a given time, i.e. only about 50, 60 or 75 percent of all the catalyst ultimately employed is utilized at a given time. Figure 6 shows that after a period of operation, the first catalyst bed in a system becomes inert and performs no useful function, requiring the succeeding catalyst bed to perform an increased hydrosulfurization load. When the metals concentration, i.e. nickel and vanadium, on the catalyst in the first catalyst bed reaches a certain level, the catalyst becomes dead for all practical purposes for desulfurization. In this condition, the catalyst performs no useful function but becomes a detriment since it contributes to the pressure drop through the reactor system.

By the method of the present invention the amount of catalyst permitted to become dead is controlled and the dead catalyst is replaced by fresh catalyst according to a prescribed schedule so that the reactor system tends to contain primarily fresh, active catalyst. If the dead catalyst is not removed, the reactor must be built large enough to store the large amount of dead catalyst that has accumulated over the entire cycle life so that the reactor is larger than it needs to be.

The removal of dead catalyst from the reactor system and replacing it with fresh catalyst is accomplished by utilizing at least one guard chamber means and preferably two swing reactors at the front end of the main hydrodesulfurization reactor. The general configuration of a process utilizing swing reactors is shown in FIG. 9. Only one swing reactor is used for processing at a time permitting the catalyst in the other one to be changed. Then, when the catalyst in the on-stream reactor is saturated, the feed stream is switched to the other reactor, which contains fresh catalyst; and this cycle is repeated until the catalyst in the main reactor is no longer usable. The whole unit is then shut down and all of the catalyst is replaced.

In accordance with this invention a method has been developed for calculating the minimum percentage of the total catalyst volume that must be on stream at a given time when utilizing a swing reactor guard chamber system. Reduction of the instantaneous on-stream quantity of catalyst in accordance with the present invention not only reduces the pressure drop in the system but also permits a reduction in reactor size which represents a considerable economic advantage. The reduction in reac tor size is a function of the size of the swing reactors, the number of times the catalyst is changed, the operating conditions and the process kinetics. The present invention differs appreciably from the concept of merely changing the catalyst at the front part of the reactor system halfway through a reaction cycle. The following calculation shows that reactor size and the instantaneous amount of on-stream catalyst can be minimized in accordance with the present invention.

Some important variables in the design of swing-reactor operation are defined below:

L-Assumed or desired cycle life of catalyst.

Y-Total weight of catalyst required including that in the main reactor and the total of all the catalyst changes.

XInstantaneous weight of catalyst in reaction zone including the main reactor plus one swing reactor.

otFraction of catalyst in reaction zone in swing reactor; (aX) is the amount of catalyst in the swing reactor, and X (la) that in the main reactor.

n--Number of times that swing reactor is switched; (n+1) is the total number of catalyst batches in the swing-reactor part of the reaction zone.

The desired cycle life L is assumed; for example, L can be 10 or 11 months for a hydrodesulfurization process. The total weight of catalyst Y required to achieve this cycle life is determined by either experiment or a detailed process simulation based upon a fundamental mathematical model. The quantities L and Y are, in general, fixed. The problem is to determine X, a, and n for a given L and Y.

The basic equation describing the relation between the variables of a swing-reactor operation is:

This equation says that the difference between the total amount of catalyst to be used and the amount of catalyst in the reaction zone at a given time is equal to the number of switches times the amount of catalyst switched at any one time. The main reasons for using swing reactors is to reduce the instantaneous amount of catalyst and the size of the reactor. Thus, the objective of the mathematical analysis is to develop a minimum value for the amount of catalyst in the reaction zone X.

To minimize X, the catalyst in the swing reactor must be switched as fast as possible but before it can be switched, the catalyst must be deactivated. Thus, the catalyst must be used as fast as possible and for the hydrodesulfurization process this means that the catalyst must be saturated with metals as fast as possible. Therefore, the operating conditions must be such that the reaction proceeds at the maximum allowable rate, i.e. the swing reactor should be operated constantly at the maximum allowable temperature subject to the constraints imposed by mechanical and metallurgical limitations, a tolerable amount of hydrocracking, and a limit on the total amount of quench between swing and main reactors.

The length of time t that any particular batch of catalyst is in the swing reactor is equal to the cycle life divided by the number of batches:

In this length of time, each batch of catalyst becomes saturated or dead; and this length of time to become deactivated is a function of the operating temperature T and the amount of catalyst in the swing reactor (otX). This function is defined by the process kinetics. For the hydrodesulfurization process, this saturation function is linear with the amount of catalyst (aX), and a sample saturation function is shown in FIG. 10. The saturation function is described by the equation:

t=b[T]+k(c\eX) (3) where the intercept b is a function of the operating temperature T, and m is 1.

Equations 1, 2 and 3 can be used to demonstrate how the overall reactor size can be reduced by switching batches of catalyst into and out of the reactor system with the swing reactors. The reduction in reactor size can be calculated for different combinations of number of switches 11, reactor size X, and fractional amount of catalyst in the swing reactor 01. The procedure to use for the calculations for a given process is:

Step 1: Assume a number of switches 11.

Step 2: Calculate batch time t from Equation 2, using the given cycle life L.

Step 3: Calculate the amount of catalyst in the swing reactor (11X) from the saturation function at the assumed operating temperature T, Equation 3, or from FIG. 11 which is a graphical representation of Equation 3.

Step 4: Calculate reactor size X from Equation 1 with n,

(aX), and the given total amount of catalyst Y. Step 5: Calculate a from (04X) and X.

From these calculations a graph of reactor size X vs. number of switches 11 can be prepared. In general, there will be an optimum number of switches that will produce a minimum reactor size. The optimum number of switches and the corresponding values of X and a can be predicted mathematically, as shown below.

The three pertinent equations for a swing-reactor operation are:

where the variables have the same definitions as indicated above and the saturation function, Equation 3, has been generalized. Substituting (2) and (3) in to (1) gives:

Equation 7 predicts the value of (04X) that will produce the maximum reduction in reactor size. Because the intercept b and the slope k are functions of the operating temperature and the metals saturation level chosen for the swing reactor, the optimum value of (aX) is a function of these quantities also.

The optimum value of (04X) can be calculated from Equation 7. Then t is calculated from Equation 3. Then 11 is calculated from Equation 2. The number of switches must be an integer. This constraint is not included in the above calculation so that the value of 11 found at the optimum may not be an integer. In this case, the optimum number of switches will be the integer closest to the optimum.

To provide a typical calculation, a process having the following operating conditions may be considered.

Feed: 50 percent Kuwait reduced crude Catalyst: M12 inch NiCoMo on alumina extrudates in four equal beds, the first bed being a swing guard chamber LVHSV: 0.55 volume feed per hour per volume catalyst LWHSV: 0.79 weight feed per hour per weight catalyst Overall space time: 1.266 weight catalyst-hour per weight feed H partial pressure: 1,830 p.s.i.

The calculation method can be applied to processes other than hydrodesulfurization wherein the catalyst is subject to progressive saturation with a deactivative material by determining the kinetics of the process to determine the relationship between t and (aX). Although the constant 111 is equal to 1 for the hydrodesulfurization process, in other processes m can be a constant other than 1. The kinetics for a hydrodesulfurization process having the above conditions yields the plot of 1 vs. (11X) for various temperatures and degrees of catalyst saturation with metals as shown in FIG. 11. FIG. 11 is a graphical presentation of the equation:

for the hydrodesulfurization process and the values of t and (11X) in the equation can be obtained directly from FIG. 11.

FIG. 12 represents a hydrodesulfurization process with no guard chamber switching in which reactor temperature, catalyst age and sulfur level are shown. Reaction conditions are those listed above. FIG. 12 shows that to obtain a bottoms product sulfur level of 1 percent there is a catalyst age limitation of 353 days and a reactor temperature limitation of 790 F. Criteria other than product sulfur level can also determine process temperature limitations, such as hydrocarbon cracking and metallurgical limitations in the reactor.

Calculations were made assuming 0, 1, 2, and 3 switchings of the swing reactors. The swing reactors Were each run to percent saturation with metals and with a constant average temperature of 750 F. The space times for the main reactor and the swing reactors were calculated, for a given 11, from Equations 1, 2, and 3. Detailed process simulations with a fundamental mathematical model showed that the life was about the same for all the runs, as the theory requires.

For one switching, the space time in the swing reactor was calculated to be 0.468 with an overall space time in the reactor system of 0.798. The calculation for one switching is achieved by following Step 1 through 5, above, as follows:

Step 1: 11:1 for 1 switching Step 2: From Equation 2, L is limited to 353 days (FIG.

12), so that t= g 176 days Step 3: From Equation 3 or FIG. 11 which is a graphical representation of Equation 3, at 176 days and 750 F. (aX) =0.468 pound of catalyst hour per pound of feed.

Step 4: From Equation 1,

For one switching, the space time in the swing reactor was calculated to be 0.468 with an overall space time in the reactor system of 0.798. For simulation purposes, the main reactor, constituting the difference between 0.798

and 0.468 or 0.330, was divided into three equal sections of 0.110 each so that the reactor system simulated had a total of four sections. Thus, the required reactor size was reduced to 100 0.798/1.266=63% of the base case. The first bed was switched at 176 days.

For two switchings, the calculated space times were 0.230, 0.192, 0.192 and 0.192 for reactor 1 (the swing reactor), 2, 3 and 4, respectively. The required reactor size was 63.5% of the base case, slightly higher than with one switching. Thus, a minimum required space time (or reactor size) was achieved with one switching.

For three switchings, the overall space time was 0.933 with the beds distributed 0.111, 0.274, 0.274 and 0.274. The reactor size was 73.8% of the base case.

FIG. 13 is a plot of the required reactor sizes vs. the number of switchings and shows that the optimum number of switchings is one for an operating temperature of 750 F. and 95% saturation. FIG. 14 is the plot for an operating temperature of 775 F. The optimum number of switchings for this case is two with an increased reduction in. reactor size. The plot shows that the required reactor size for these conditions is approximately one-half that with no catalyst changes. At metals loadings lower than 95% even smaller reactors will be optimum, although possibly more catalyst would be required to obtain the same cycle life.

EXAMPLE 1 A series of tests were conducted to illustrate the temperature advantage of a small particle size catalyst of the present invention. These tests were conducted by employing NiCoMo on alumina catalysts of various sizes for hydrodesulfurizing a 36 percent Kuwait reduced crude from which furnace oil having an 800 F. TBP had been removed at a 2000 p.s.i.a. partial pressure of hydrogen and a space velocity of 3.0 liquid volumes per hour per volume of catalyst. The charge was 78 percent desulfurized to a 1.0 percent product sulfur content. The arrangement of the reactor was such that there was no significant or readily detectable pressure drop in any of the tests. FIG. 1 shows the eifect of catalyst size upon the initial temperature required to produce a product containing 1 percent sulfur. The solid line is based upon initial temperatures determined in tests with /s inch and inch diameter extrudate catalyst whose particle size is above the range of this invention. The dashed extrapolation of the solid line indicates that a inch diameter extrudate catalyst would be expected to require an initial temperature of about 775 F. However, FIG. 1 surprisingly shows that the 1 inch diameter extrudate catalyst requires an initial temperature of only 750 F. It is noted that the surface area defined by the pores of all three catalysts tested is about the same. The position of the data point for the 3 inch catalyst is highly surprising because if the dashed line in FIG. 1 were curved downwardly towards the inch catalyst data point the resulting curve would tend to indicate that as catalyst particle size becomes very small the activity of the catalyst becomes unlimited, which is obviously unreasonable. Therefore, the straight configuration of the dashed extension of the curve in FIG. 1 is a reasonable extrapolation of the solid line and the position of the inch data point is highly unexpected.

EXAMPLE 2 When catalysts similar to the inch catalyst of Example 1 except that the particle size is smaller within the range of this invention, such a or inch, or except that the particle size is larger within the range of this invention, such as or inch, are utilized under the conditions of Example 1, the initial temperature in each instance to achieve hydrodesulfurization to one percent sulfur is at about the same level as that shown in FIG. 1 for the & inch catalyst.

EXAMPLE 3 When catalyst compositions other than NiCoMo on alumina, such as NiCoMo on silica alumina, CoMo on 10 alumina, NiW on alumina, NiW on silica alumina, NiW on silica magnesia or NiMo on alumina having a particle size within the range of this invention are utilized under the conditions of Example 1 to achieve hydrodesulfurization to one percent sulfur, a similar unexpected temperature advantage is realized as compared to the extrapolated temperature based on larger size particles of the same composition.

EXAMPLE 4 Tests were made which demonstrate that a inch nickel-cobalt-molybdenum on alumina extrudate is not only capable of hydrodesulfurizing a reduced crude oil to a one percent sulfur level at a considerably lower initial temperature than a similar catalyst in the form of a ,5 inch extrudate but also is capable of maintaining a lower hydrodesulfurization temperature with age. The tests with the ,4, inch catalyst were based on a 0.55 liquid hourly space velocity and a hydrogen partial pressure of 1830 pounds per square inch. The reactor pressure drop was about 50 pounds per square inch. The charge was a 50 percent Kuwait reduced crude. The reaction was performed in a single reactor having three separate beds and recycle hydrogen gas was used as a quench after each bed. There was no separate guard chamber before the reactor. The first, second and third beds contained 13.3 percent, 41.6 percent and 45.1 percent of the total catalyst, respectively. Typical data for the test utilizing the $4 inch catalyst are shown below and the general data is illustrated in FIGS. 5 and 6. FIG. 5 shows the aging characteristics for the entire & inch catalyst reactor as compared with a comparable aging run with a ,5 inch catalyst reactor. FIG. 6 shows the aging characteristics for the individual beds within the inch reactor and shows that when the first bed becomes deactivated the second bed assumes a greater desulfurization load.

The test utilizing the inch nickel-cobalt-molybdenum on alumina catalyst was performed at a 1.1 liquid hourly space velocity, but is illustrated in FIG. 5 on a basis comparable to the 0.55 space velocity of the test made with inch catalyst. The total pressure for the inch catalyst test was 2500 pounds per square inch gauge. 5000 s.c.f./bbl. of gas was charged to the reactor. The reactor contained four catalyst beds and recycle gas was used as a quench after each bed. The average reactor temperature was increased throughout the test to maintain a 1.0 percent by weight sulfur level in the 660 F.+residual product. Typical data for both the inch and the 34 inch catalyst tests are shown below.

142 inch /ifl inch catalyst catalyst Oil charge 1 1 Catalyst..... E Volume, cc 2, 294 2 254 Weight, gram 1, 54s. 0 1 its. 0 Age, days this measuremen 97. 6 l 87 6 Total throughput, volume 0 p volume catalyst 1, 293 2 323 Operatmg conditions:

Reactor bed temp., F. (Inlet, Outlet) 694, 716 Reactor pressure 2 050 2 510 Avg. reactor temp., F. 703 Y 784 Space yellocityz o r. vol 0.54 1. 11 Wt./hr./Wt 0.78 1. 36 Reagtorflchargc:

.c. Percent H 2% Makseupf .c.

Percent HL 3 3 Recgclefgfisil .c. Percent H2 3, 3 33 Product yields, wt. percent:

Bottoms (660 F. 91. 1 84. 7 Furnace oil (380-660 F.) 4. 9 9. 4 Naphtha (I.B.P. 380 F.) 0.8 2. 2 Gas 5. 4 5. 4 Chemical hydrogen consnmp 11, b1 476 617 Hydrogen sulfide, S.c.f./bbl 139 127 The charge and product inspections for the test employing the inch catalyst are as follows:

Product Charge bottoms Gravity, API 14. 6 20. 1 Sulfur, percent by wt 4. 07 1.03 Nitrogen, percent by wt 0.22 0. 17 Carbon residue, percent by w 8. 50 4.597 91 3 19, 094

No'rn.Charge, cracked at 995.

The 660 F.+residual oil product inspections for the inch catalyst test are as follows:

EXAMPLE Tests were made to show the effect of temperature upon liquid yield in a hydrodesulfurization process. The tests were made in a pilot plant equipped with a four-bed 2254 cc. adiabatic reactor. Reactor charge gas was used as quench'between the catalyst beds for temperature control. The charge stock passed through a cotton fiber cartridge filter before it was preheated and charged to the reactor. The filter which was at steam tracing temperature takes out most of any solid contaminants in the feed, but does very little in removing any small or organically combined metals present in the charge stock.

The reactor effluent flowed into a high pressure separator where hydrogen-rich gas was separated from the hydrocarbon liquid. The hydrogen-rich gas was scrubbed with 3 percent to 5 percent diethanolamine and water and recycled to the reactor. After high pressure separation of high pressure hydrogen-containing gas, the liquid product flowed to distillation towers where gases, naphtha, furnace oil, and residual were removed from the unit.

The charge to the unit was a 50 percent Kuwait reduced crude. The operation was designed to produce a 660 F.+ product having a 1.0 percent sulfur level. The catalyst was inch extruded NiCoMo on alumina. The operating conditions were 2500 pounds per square inch gauge total pressure, 1.1 liquid hourly space velocity and 5000 s.c.f./bbl. of 80 percent hydrogen with recycle gas quench as required for temperature control. The results of the tests are illustrated in FIG. 4 and in the following data.

Catalyst.NiCoMo on alumina having 0.97 wt. percent cobalt, 8.6 wt. percent molybdenum and 0.59 wt. percent nickel Age:

Days at this measurement 45.9 BbL/lb. 4.43 Space velocity: LHSV 1.1 Average reactor temp F. 760

12 Reactor gas:

InletS.c.f./bbl. 5008 Percent H 82 QuenchS.c.f./bbl. 2920 5 Percent H 82 Reactor pressure: p.s.i.g 2500 Hydrogen consumption: S.c.f./bbl. 623 Product yields: Percent by weight H S 3.4 C 0.2 C 0.1 C 0.2 Ca, 0.2 C -380 F. 1.5 380-460 F. 1.4 460-600 F. 2.8 600-660 F. 2.5 660 F.+ 88 6 2O EXAMPLE 6 Tests were conducted to illustrate the effect of a change in hydrogen partial pressure upon the temperature required to hydrodesulfurize a reduced crude to 1 percent sulfur in the residual product. The comparative tests were performed by, in one case, not recycling hydrogen containing light hydrocarbons which build up in the hydrogen stream and reduce the partial pressure of the hydrogen in the stream but instead charging to the hydrodesulfurizer only fresh hydrogen charge having a uniform hydrogen purity. In the other case, a recycle hydrogen stream which was not subjected to naphtha scrubbing to remove light hydrocarbons so that the hydrogen partial pressure therein continually decreased throughout the test was recycled to the hydrodesulfurizer. The reactor system catalyst and operating conditions for both tests are generally the same as that described in the tests of Example 4. The results are illustrated in FIG. 3. The solid line in FIG. 3 represents the test utilizing only fresh hydrogen at 1830-1850 pounds per square inch of hydrogen pressure. The broken line in FIG. 3 represents the test wherein non-naphtha scrubbed recycle gas is recycled causing hydrogen partial pressure to continually decrease so that at the last data point shown the hydrogen partial pressure was 1720-1740 pounds per square inch. The following data are representative of the test represented by the broken line.

Oil charge, percent Kuwait reduced crude 50 Catalyst-4& inch NiCoMo on alumina:

Volume, cc. 2296 Weight, gram 1771 Age, days at time of this measurement 7.2 Throughput, vol. oil per vol. cat. 96

Reactor bed temp. F.:

(Inlet, Outlet) 668, 690 Operating conditions:

Reactor pressure, p.s.i.g 2058 Avg. reactor temp, F. 676 Space velocity- Vol./hr./v0l. 0.53 Wt./hr./wt. 0.66 Reactor gas charge- S.c.f./bbl. 4462 percent H 88 Makeup gas- S.c.f./bbl. 587

Percent H 94 Recycle gas- S.c.f./bbl 3874 Percent H Product yields, wt. percent:

Stripper bottoms 92.5 Furnace oil 4.7

Naphtha 0.6 Gas 3.7 Net hydrogen sulfide, S.c.f./bbl 108 Product Charge stripper o ottoms Gravity, API 15. 7 20. 6 V1scosity, SUS D2161:

100 F 4, 006 2, 181 210 F 1. 8 114. 8 Carbon, wt. percent 84. 52 85. 2 Hydrogen, wt. percent 11.43 11.68 N ltrogcn, wt. percent 0. 20 0. 17 Sulfur, wt. percent 4. 06 1. 11 Carbon residue, wt. percen 8. 5. 12 Nickel, p.p.m 16 4. 7 Vanadium, ppm 54 6- 1 Heat of combustion, Btu/lb 1. 18, 423 19, 908 Distillation, vacuum, F.:

t 5%.--- 608 654 At 10 674 682 At 20 g 762 750 At 30% 829 807 At 40% 888 866 At 50%- 925 At 60% 992 NoTr:.Charge oil, cracked at 888. Product stripper bottoms cracked at 1,011.

EXAMPLE 7 Simulation experiments were conducted to show the effect of catalyst particle size on pressure drop in hydrodesulfurization process in reactors of various diameters. All tests were made with the same liquid hourly space velocity in a single bed reactor, charging a 75 percent reduced Kuwait crude, using recycle hydrogen and maintaining hydrogen purity of 77 percent, using reactor inlet and outlet temperatures of 780 and 815 F., respectively, a reactor inlet pressure of 2500 p.s.i.-g., and a 1.0 liquid hourly space velocity. Three series of tests were made utilizing reactors of various diameters with inch, inch and & inch NiCoMo on alumina catalyst particles. The results are illustrated in FIG. 2.

The following description of a process of this invention is made in reference to FIG. 7. FIG. 7 itself indicates suitable temperature and pressure conditions at various points in the process, and these indicated conditions are not reiterated in the following description.

Referring to FIG. 7, a full crude or a reduced crude, such as a 50 percent reduced Kuwait crude, which contains the entire asphaltene content of the full crude and therefore also contains substantially all of the nickel, vanadium and sulfur content of the full crude, is charged to the process through line 10 and is pumped by pump 12 through line 14, preheater 16, line 18, solids filter 20 and line 22 to drum 24. From drum 24 the liquid oil charge is passed through line 26 to feed pump 30. Liquid from pump 30 is admixed with hydrogen from line 52 and passed through line 32, valve 34, preheater 36 line 38 and furnace 40.

A mixture of fresh and recycle hydrogen is introduced into the liquid charged to the reactor prior to preheating thereof. Recycle hydrogen is admitted to the liquid charge through line 42 and valve 44. Makeup hydrogen is charged through line 46, compressor 48 and valve 50. A mixture of fresh and recycle hydrogen is introduced to the relatively cool liquid charge through line 52.

The preheated mixture of liquid charge and hydrogen is charged through line 54 to either guard chamber 56 containing a bed of catalyst 58 or guard chamber 57 containing bed of catalyst 59. Each guard chamber is blocked off by valves as shown so that when the catalyst in one becomes dead it can be taken off-stream and filled with fresh catalyst while the other is put on stream. An efliuent stream from a guard reactor is charged to main reactor 60 containing catalyst beds 62, 64, and 66. The catalyst beds each contain NiCoMo on alumina catalyst as inch extrudates which have the following typical specifications:

Surface area-150 meters /gram Pore volume 50300 A. radius-60% to 90% of total pore volume Pore volume0.5 to 0.8 cc./ gram 14 Compacted density-O.45 to 0.65 gram/cc. Specific volume of pores-3040 cc./10O cc.

Each succeeding catalyst bed has a larger volume than the bed just prior to it. If desired there can be four, five, six or more catalyst beds in the reactor. Also, if desired, each reactor bed can have 25 percent, 50 percent, percent, or more, catalyst than the bed just prior to it.

The effluent from a guard reactor is charged through line 68 to which a temperature quenching hydrogen stream is charged through line 70 and valve 72 so that a quenched hydrocarbon and hydrogen stream is charged to the top of the reactor through line 74. The reaction stream is passed through catalyst bed 62 and due to the exothermic nature of the hydrodesulfurization reaction it is heated in passage therethrough and is quenched by recycle hydrogen entering through line 76, valve 78, and sparger 80. The cooled reaction stream is then passed through catalyst bed 64 wherein it increases in temperature so that it is then cooled by a quenching hydrogen stream entering through line 82, valve 84, and sparger 86. Temperatures between the various catalyst beds are controlled by regulating the valves in the various hydrogen quench lines to apportion recycle hydrogen flow. Finally the reaction mixture passes through catalyst bed 66 and then leaves the reactor in a desulfurized condition through line 88. The reactor eflluent stream then gives up some heat at charge preheater 36 and passes through'line 90 to air cooler 92 and valve 94. Reactor eflluent enters flash chamber 96 whence desulfurized liquid is discharged through line 98 to a distillation column 102. A gaseous stream comprising primarily hydrogen together with ammonia and hydrogen sulfide formed by nitrogen and sulfur removal from the charge and light hydrocarbons formed by thermal hydrocracking of some of the charge is discharged from flash chamber 96 through line 99.

The gaseous effluent stream passes through unit 106 to which water is added through line 108 and from which aqueous ammonia is removed through line 110. The gaseous eflluent from unit 106 is passed through line 112 to light hydrocarbon wash unit 114 to which naphtha from distillation column 102 is pumped in order to wash light hydrocarbons from the hydrogen stream. The wash naphtha is removed through line 116 and passed to flash chamber 118 Where some of the dissolved hydrocarbons are flashed off through line 120. Then the naphtha is passed through line 122, heater 124 and hot naphtha flash chamber 126 from which additional light hydrocarbons are flashed through line 128. The regenerated naphtha is recycled through line 130 with makeup naphtha entering through line 132.

The hydrogen stream is then passed through line 134 to hydrogen sulfide removal unit 136 to which an amine such as monoethanolamine is added through line 138. Hydrogen sulfide-rich amine is removed through line 140 and passed to amine regenerator 142 from which hydrogen sulfide is discharged through line 144 and recycle amine exits through line 146. Makeup amine is charged through line 148. A recycle hydrogen stream is then re turned to the reactor through line 150.

It is important to remove a substantial amount of the ammonia, hydrogen sulfide and light hydrocarbons from the hydrogen stream prior to recycle because these gases reduce the partial pressure of hydrogen in the reactor. It is the partial pressure of hydrogen rather than the total pressure in the reactor which affects hydrodesulfurization activity. It is not possible to arbitrarily increase total hydrogen pressure in the reactor to compensate for a low hydrogen partial pressure because of severe design pressure limitations in the reactor, as explained above. The recycle hydrogen is passed through compressor 154 for increasing the pressure of the hydrogen stream to the reactor.

Hydrodesulfurized residual oil is removed from dis- 15 tillation column 102 through line 156 and is used to supply heat to inlet crude oil heat exchanger 16 prior to its discharge from the system through line 158. Hydrodesulfurized furnace oil is removed from the distillation column through line 160 while naphtha product is removed from the system through line 162.

FIG. 8 is a cutaway segment of a multibed reactor in which the lower two catalyst beds are shown. FIG. 8 shows that a bed of the small size catalyst of the present invention is prepared in a manner to secure the particles from excessive relative movement and to prevent the particles from producing fines and plugging screens, both of which conditions would greatly increase the pressure drop in the reactor and tend to further nullify the temperature advantage of the small size catalyst particles.

FIG. 8 shows a steel reactor wall 1000 which may be 7 to 10 inches thick. One bed of catalyst lies above hydrogen quench line 1002 and another catalyst bed lies below this line, each bed filling the entire cross-section of the reactor. The greatest volume of the upper bed comprises inch catalyst bed 1004 which rests upon a smaller bed Oi /12 inch catalyst 1005 and inch aluminum balls 1006 which in turn rests upon a bed of /2 inch aluminum balls 1008. Beds 1005, 1006 and 1008 prevent the & inch catalyst particles from surrounding and plugging the sparger openings of hydrogen quench line 1002. Above catalyst bed 1004 is a bed 1010 of inch aluminum balls and a bed 1012 of /2 inch aluminum balls. These latter two beds contribute stabilizing weight to the 1 inch catalyst bed to prevent shifting of particles therein during flow of reactants, which shifting would tend to cause catalyst disintegration and fines formation and thereby greatly enhance pressure drop through the inch bed.

The lower bed of catalyst rests upon screen 1014. Screen 1014 is protected from plugging by inch catalyst bed 1016 by a gradual increase in particle size between it and the inch catalyst bed as indicated by inch catalyst bed 1018, /4 inch aluminum balls bed 1020 and /2 inch aluminum balls bed 1022. Proper distribution of hydrogen and liquid reactant as they approach lower catalyst bed 1016 is insured by bed 1024 of 4 inch aluminum balls and bed 1026 of /2 inch aluminum balls.

FIG. 8 shows that an elaborate arrangement is required in preparing a catalyst bed of the present invention in order that nearly all the pressure drop through the bed can be confined to the inch catalyst beds themselves with very little pressure drop at retaining screens and with a minimum of pressure increase occurring due to fines formation during the reaction.

We claim:

1. A process for the hydrodesulfurization of a crude oil or a reduced crude containing the asphaltene fraction of the crude wherein a mixture of hydrogen and said hydrocarbon oil is passed through a swing guard reactor and a main reactor in series, each of said reactors containing a proportionate amount of catalyst particles comprising Group VI metal and Group VIII metal on alumina as a compact bed of particles which are between about and ,4 inch in diameter, the hydrogen partial pressure in said process being between 1,000 and 5,000 p.s.i., applying a hydrogen quench to said main reactor and establishing the proportionate amount of catalyst in said main reactor and said guard reactor and the switching frequency of said guard reactor by the following equations 16 where:

t is the time a batch of catalyst is in the swing reactor L is the desired cycle life of the catalyst Y is the total weight of the catalyst including that in the main reactor and the total of all the catalyst changes X is the instantaneous weight of catalyst in the main reactor plus one swing reactor at is the fraction of catalyst in the swing reactor;

(11X) is the amount of catalyst in the swing reactor 11 is the number of times that swing reactor is switched;

(n+1) is the total number of catalyst batches in the swing reactor, and

b and k are constants determined by the kinetics of the hydrodesulfurization reaction at the temperature of operation.

2. The process of claim 1 wherein said catalyst particles are between and 1 inch in diameter.

3. The process of claim 1 wherein said catalyst particles are between and &4 inch in diameter.

4. The process of claim 1 wherein said catalyst comprises nickel-cobalt-molybdenum on alumina.

5. The process of claim 1 wherein the charge oil contains between about 0.002 and 0.03 weight percent of nickel and vanadium.

6. The process of claim 1 wherein the main reactor is divided into a series of separate beds and the temperature is controlled by applying a hydrogen quench between the beds.

7. A process for the hydrodesulfurization of a crude oil or a reduced crude oil containing the asphaltene fraction of the crude wherein a mixture of hydrogen and said hydrocarbon oil is passed through a swing guard reactor and a main reactor each containing a proportionate amount of catalyst particles comprising a Group VI metal and Group VIII metal on alumina as a compact bed of particles which are between about 1 and 4 inch in diameter, the hydrogen pressure in said process being between 1,000 and 5,000 p.s.i. and establishing the proportionate amount of catalyst in said main reactor and said guard reactor by the following equation:

a is the fraction of the catalyst in the swing reactor X is the instantaneous weight of catalyst in the main reactor plus one swing reactor ocX is the amount of catalyst in the swing reactor L is the desired cycle life of the catalyst, and

b and k are constants determined by the kinetics of the hydrodesulfurization reaction at the temperature of operation.

8. The process of claim 7 wherein the catalyst particles are between & and A inch in diameter.

9. The process of claim 7 wherein the catalyst particles are between and inch in diameter.

10. The process of claim 7 wherein the catalyst comprises nickel-cobalt-molybdenum on alumina.

11. The process of claim 7 wherein the charge oil contains between about 0.002 and 0.03 Weight percent of nickel and vanadium.

12. The process of claim 1 wherein the hydrogen partial pressure is 1,500 to 2,500 p.s.i.

13. The process of claim 7 wherein the hydrogen partial pressure is 1,500 to 2,500 p.s.i.

References Cited UNITED STATES PATENTS 2,963,425 12/1960 Hansen, Jr. 208216 3,072,564 1/1963 Stewart, Jr. et al. 208210 3,278,420 10/1966 Jaeger 208216 3,349,027 IO/1967 Carr et a1 2082l6 3,418,234 12/1968 Chervenak et a1. 208210 DELBERT E. GANTZ, Primary Examiner G. J. CRASANAKIS, Assistant Examiner 

